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Integrated Masters in Chemical Engineering

Packed Bed Membrane Reactor for the Water-Gas Shift Reaction: Experimental and Modeling

Work

Masters Thesis

of

Joel Alexandre Moreira da Silva

Developed in the framework of the course of dissertation

performed at

Multiphase Reactors Group, Department of Chemical Engineering & Chemistry,

Eindhoven University of Technology

Supervisor at TU/e: Eng. Arash Helmi, Dr. Fausto Gallucci and Dr. Martin Van Sint Annaland

Chemical Engineering Department

July of 2013

Packed Bed Membrane Reactor for the Water-Gas Shift Reaction: Experimental and Modeling Work

i

Acknowledgements

I would like to deeply thank Engineer Arash Helmi for all the advice, patiente,

support, encouragement and dedication to this project. His help was fundamental to this

project.

I would also like to thank Doctor Fausto Gallucci for all the good advices given and all

the help provided during this project.

I would also like to thank Doctor Martin Van Sint Annaland for all the suggestions made

regarding the scope of my project and for making this project possible to happen.

Also a special thanks to Masters student Rogelio Gonzlez for all the help while using

the phenomenological models that he was developing.

Also a very special thanks to Joris for the technical support provided to the set-up I

used and for all the technical advices.

I would like to thank all SMR group members for making this group such a good one,

with such a good environment.

Finally I would like to thank my family and friends for their encouragement and

support. They played undoubtedly a very important role during this dissertation project.


ii

Abstract

The main goal of this work was to perform the experimental study of a water-gas shift

(WGS) reaction and produced hydrogen separation unit using a Pd membrane. By combining

both elements in a single unit it is expected to obtain conversions that go beyond the

thermodynamic equilibrium, for a specific operating temperature and feed stream

composition, due to the selective removal of hydrogen in the reaction media. It is also

expected to obtain a highly pure hydrogen stream that can be used in systems that are highly

sensitive to CO poisoning for example. The second main goal of this work was to validate both

1D and 2D pseudo-homogeneous models for the permeation of hydrogen through the Pd

membrane and later for the WGS reaction inside the packed bed membrane reactor (PBMR).

The possible demonstration of the existence of the concentration polarization effect,

considered in the 2D model, was one of the aspects of more focus.

A lab set-up that is continuously available for permeation tests was used. In a first

stage this set-up was used to perform the characterization of the permeation of hydrogen

through several Pd membranes and in a later stage the set-up was modified so that it would

be possible to study the performance of the WGS reaction inside a PBMR. Regarding the

catalyst characterization, since it had been already done by a PhD student it wasnt part of

the scope of this project. Regarding the activation of the catalyst used, which in this case was

0.5Pt/6CeTiO2, it was performed in a lab set-up that is continuously available for kinetics

tests. Finally both 1D and 2D models, developed in Delphi 7, which simulate the WGS reaction

inside a PBMR were used to simulate the permeation of hydrogen through the Pd membrane

used in the membrane reactor. The comparison between the results obtained using both

models and the experimental results as well as the quantification of the concentration

polarization effect were done.

It can be concluded that the Pd-Ag membrane reactor allowed the conversion of CO

beyond the thermodynamic equilibrium, as expected, and simultaneously produce a highly

pure hydrogen stream that meets the requirements of the polimer electrolyte membrane fuel

cells. It was also possible to validate both phenomenological models for the permeation of

hydrogen through the Pd-Ag membrane and verify that the concentration polarization effect

is not negligible for some of the conditions tested.

Keywords: hydrogen; water-gas shift; packed bed membrane reactor; concentration

polarization.


iii

Resumo

O principal objetivo deste trabalho foi efetuar o estudo experimental de uma unidade

de reao de gs-de-gua (water-gas shift, WGS) e de separao do hidrognio produzido por

via de uma membrane de Pd. Ao combinar ambos os elementos numa nica unidade espera-se

conseguir obter converses que vo alm do equilbrio termodinmico, para uma determinada

temperatura de operao e composio da corrente de alimentao, devido remoo

seletiva do hidrognio presente no meio reacional. Espera-se tambm a obteno de uma

corrente de hidrognio de elevada pureza que possa ser utilizada em sistemas altamente

sensveis ao envenenamento por CO por exemplo. O segundo grande objetivo deste trabalho

foi a validao de ambos os modelos pseudo-homogneos 1D e 2D para a permeao de

hidrognio atravs da membrana de Pd e posteriormente para a reao de WGS no interior de

um reator de membrana de leito fixo. A possvel demonstrao da existncia do efeito da

polarizao da concentrao, tido em conta pelo modelo 2D, foi um dos aspetos ao qual foi

dada mais ateno.

Para isto foi utilizada uma instalao laboratorial que est continuamente disponvel

para testes de permeao. Numa primeira fase esta instalao foi utilizada para efetuar a

caracterizao da permeao de hidrognio atravs de vrias membranas de Pd e numa fase

posterior a instalao foi alterada para que se pudesse estudar a reao de WGS no interior

de um reator de membrana de leito fixo. Em relao caracterizao do catalisador, esta j

tinha sido feita por um aluno de PhD sendo que no fez parte da ordem de trabalhos deste

projeto. Relativamente ativao do catalisador usado, que neste caso foi 0.5Pt/6CeTiO2,

esta foi efetuada numa instalao laboratorial que est continuamente disponvel para testes

cinticos. Por fim os modelos 1D e 2D, desenvolvidos no software Delphi 7, que simulam a

reao de WGS no interior de um reator de membrana de leito fixo foram usados para simular

a permeao de hidrognio atravs da membrana de Pd usada no reator de membrana. A

comparao entre os resultados obtidos usando os modelos e os resultados experimentais foi

feita e a quantificao do efeito da polarizao da concentrao foi efetuada.

Concluu-se que o reator de membrana de Pd-Ag permite converter CO para alm do

equilbrio termodinmico, tal como era suposto, e produzir em simultneo uma corrente de

hidrognio altamente puro que vai de encontro ao requsitos das clulas de combstivel de

membrana eletroltica polimrica. Foi tambm possvel validar ambos os modelos

fenomenolgicos para a permeao de hidrognio atravs da membrana de Pd-Ag e verificar

que o efeito da polarizao da concentrao no desprezvel para certas condies.

Palavras Chave: hidrognio; water-gas shift; reator de membrana de leito fixo; polarizao

da concentrao.


iv

Declaration

I declare under oath that this work is original and that all non-original contributions

were adequately referenced with the reference identification.


v

Syllabus

Figures Syllabus ............................................................................................. ix

Tables Syllabus ............................................................................................ xiii

1 Introduction ............................................................................................ 1

1.1 Framework and Project Presentation ...................................................... 1

1.2 SMR Presentation ............................................................................... 3

1.3 Work Contributes ............................................................................... 3

1.4 Thesis Organization............................................................................. 3

2 Context and State of the Art ....................................................................... 5

2.1 History of Hydrogen and the WGS Reaction ............................................... 5

2.2 Thermodynamics of the WGS Reaction ..................................................... 7

2.3 Catalysts for the WGS Reaction .............................................................. 9

2.4 Mechanisms and Kinetics ..................................................................... 11

2.5 Hydrogen Purification ......................................................................... 15

2.6 Membrane Reactors ........................................................................... 17

2.7 Dense H2 Perm-Selective Membranes for Membrane Reactors ....................... 18

2.8 The WGS Reaction in Packed Bed Membranes Reactors ............................... 20

2.9 Modeling Studies on the WGS Reaction Carried in Packed Bed Membrane Reactors

22

3 Technical Description ............................................................................... 23

3.1 Membrane Characterization ................................................................. 23

3.1.1 Experimental set-up .................................................................................... 23

3.1.2 Experiments performed ................................................................................ 24

3.1.3 Results and discussion .................................................................................. 25

3.2 WGS Reaction Tests in a Packed Bed Membrane Reactor ............................. 32

3.2.1 Experimental set-up .................................................................................... 32

3.2.2 Experiments performed ................................................................................ 33

3.2.3 Results and discussion .................................................................................. 34


vi

3.2.3.1 Influence of the reaction temperature ....................................................... 34

3.2.3.2 Influence of the GHSV ........................................................................... 36

3.2.3.3 Influence of the H2O/CO ratio ................................................................. 38

3.2.3.4 Influence of the H2 partial pressure difference in the membrane ....................... 40

3.3 Validation of Both 1D and 2D Phenomenological Models .............................. 41

3.3.1 Models description ...................................................................................... 41

3.3.2 Validation of the 1D and 2D phenomenological models .......................................... 41

4 Conclusions ............................................................................................ 44

5 Overall evaluation of this work ................................................................... 46

5.1 Achieved Goals ................................................................................. 46

5.2 Limitations and Future Work ................................................................ 46

5.3 Final Appreciation ............................................................................. 47

6 References ............................................................................................ 48

Appendix A Results of the Permeation Tests for the Thin Pd Membranes ................... 51

A.1 Experiments Performed for the Thin Pd Membranes .................................... 51

A.2 Results Obtained ................................................................................. 51

Appendix B Gas Chromatograph Calibration ........................................................ 59

Appendix C Supporting Data for the Permeation Results of the Pd-Ag Membrane ......... 63

C.1 Regressions for and ............................................................ 63

C.2 Comparison Between the Predicted Flux Values and the Measured Flux Values

(Parity Plots) ............................................................................................. 66

Appendix D Catalyst Characterization ............................................................... 69

D.1 - Kinetics of the WGS Reaction on the 0.5Pt/6CeTiO2 Catalyst .......................... 69

Appendix E Verification of the Catalyst Activity .................................................. 75

Appendix F Determination of the Parameters of the Sieverts-Langmuirs Model Equation

................................................................................................................. 77

Appendix G Quantification of the Impact of Both Concentration Polarization and CO

Poisoning Effects on the Predicted H2 flux through the Pd-Ag Membrane..................... 79


vii

Notation and Glossary

Concentration of atomic hydrogen at the spacial position along the membrane thickness

molm-3

Effective diffusion coefficient of atomic hydrogen m2s-1

Diameter of the membrane tube m

Apparent activation energy kJmol-1

Apparent activation energy of the Pd membrane kJmol-1

Atomic hydrogen diffusion flux through the metal lattice molm-2s-1

Flux of H2 through the membrane molm-2s-1

Predicted total flux of H2 through the Pd-Ag membrane mol s-1

Forward reaction rate constant mol h-1Pa-2

Forward reaction rate constant mol h-1Pa-1

Reverse reaction rate constant mol h-1Pa-1

Pre-exponential factor s-1

Langmuirs adsorption constant for CO Pa-1

Equilibrium constant

Forward reaction rate constant mol h-1Pa-(a+b+c+d)

Equilibrium adsorption constant of species Pa-1 or Pa-1.5

Membrane constant molm-1s-1Pa-x

Membrane length m Molar flow of H2 in the feed stream molh

-1

Molar flow of H2 in the permeate outlet molh-1

Molar flow of H2 in the retentate outlet molh-1

Partial pressure of CO Pa

Partial pressure of H2 in the permeate side Pa

Partial pressure of H2 in the retentate side Pa

Partial pressure of component Pa Permeability of the Pd-Ag membrane to H2 molm

-1s-1Pa-1

Pre-exponential factor molm

-1s-1Pa-1

Volumetric flow of H2 at the permeate outlet mLNmin

-1

Volumetric flow of CO at the retentate inlet mLNmin

-1

Volumetric flow of CO at the retentate outlet mLNmin

-1

Volumetric flow of H2 at the retentate outlet mLNmin

-1

Experimental reaction rate mol h-1

Ideal gas constant kJK-1mol-1

H2 recovery

Forward reaction rate mol

h-1

Pre-exponential factor mol h-1

Sorption coefficient of hydrogen in the metal lattice molm-3Pa-0.5

Absolute temperature K

CO conversion

Conversion of CO at the equilibrium

Molar fraction of the component at the reactor inlet


viii

Greek letters

H2 permeance reduction factor

Approach to equilibrium Pd-Ag membrane thickness m

Reaction enthalpy at 298 K kJmol-1

Relative error

Superscripts

Forward reaction orders for H2O, CO, H2 and CO2 respectively Pressure exponent

List of acronyms

CSTR Continuous Stirred-Tank Reactor FBMR Fluidized Bed Membrane Reactor GHSV Gas Hourly Space Velocity MR Membrane Reactor PBMR Packed Bed Membrane Reactor PBR Packed Bed Reactor PEMFC Polymer Electrolyte Membrane Fuel Cell SRM Steam Reforming of Methane TR Traditional Reactor WGS Water-Gas Shift -GC Micro Gas Chromatograph


ix

Figures Syllabus

Figure 1 - Process flow in a typical fuel processor operating in the steam reforming mode. Taken from

[4] ............................................................................................................................2

Figure 2 CO equilibrium conversions of a typical reformate stream from a SRM process for different

steam to dry gas (S/G) ratios. Taken from [5] .......................................................................8

Figure 3 - Process schemes for the hydrogen production and purification; (a) traditional process

considering an absorption and catalytic approach for hydrogen purification (b) PSA-based hydrogen

purification. Taken from [5] ........................................................................................... 16

Figure 4 - Hydrogen production and purification based on the WGS MR unit. Taken from [5] ............ 17

Figure 5 Composition of the outlet stream of a MR and a traditional process with a typical

composition of syngas coming out of a reformer and a H2O/CO molar ratio of 1. Taken from [27] ...... 18

Figure 6 Scheme of a PBMR for the WGS reaction. Adapted from [10] ...................................... 20

Figure 7 - Scheme of the experimental set-up for the Pd-based membrane testing. ...................... 24

Figure 8 Hydrogen flux through the Pd-Ag membrane as a function of the difference between the

square roots of the hydrogen partial pressure in the retentate and permeate sides. ...................... 27

Figure 9 Comparison between the Arrhenius plot obtained in this work and others reported in the

literature. The line represents the regression of the data of this work with equation (20). .............. 30

Figure 10 H2 flux as a function of the CO concentration in the feed for different H2 trans-membrane

partial pressure differences at 400 C. .............................................................................. 31

Figure 11 - Scheme of the experimental set-up for the WGS reaction tests in a PBMR. ................... 32

Figure 12 - Influence of the reaction temperature on the CO conversion for the WGS reaction over the

0.5Pt/6CeTiO2 catalyst in the PBR and in the Pd-Ag PBMR. ...................................................... 35

Figure 13 - Influence of the reaction temperature on the H2 recovery for the WGS reaction over the

0.5Pt/6CeTiO2 catalyst in the Pd-Ag PBMR. ......................................................................... 35

Figure 14 - Influence of the GHSV on the CO conversion for the WGS reaction over the 0.5Pt/6CeTiO2

catalyst in the Pd-Ag PBMR. ........................................................................................... 36

Figure 15 - Influence of the GHSV on the H2 recovery for the WGS reaction over the 0.5Pt/6CeTiO2

catalyst in the Pd-Ag PBMR. ........................................................................................... 37

Figure 16 - Influence of the H2O content in the feed on the CO conversion for the WGS reaction over

the 0.5Pt/6CeTiO2 catalyst in the Pd-Ag PBMR. .................................................................... 38

Figure 17 - Influence of the H2O content in the feed on the H2 recovery for the WGS reaction over the

0.5Pt/6CeTiO2 catalyst in the Pd-Ag PBMR. ......................................................................... 39


x

Figure 18 - Comparison of the H2 permeating flux obtained using the 1D and 2D models and 1D and 2D

models with Sieverts-Langmuirs model equation with the experimental one for a H2 partial pressure in

the retentate side of 3.5 bar at 400 C.............................................................................. 43

Figure A.1 Flux of H2 through the 3.5-4.0 m thick Pd membrane as a function of the H2 trans-

membrane partial pressure difference for different H2 compositions at 400 C. ............................ 51

Figure A.2 - Flux of H2 through the 4.0-5.0 m thick Pd membrane as a function of the H2 trans-

membrane partial pressure difference at 300 C. ................................................................. 54

Figure A.3 - Linear regression of the H2 flux through the 3.5-4.0 m thick Pd membrane as a function

of the difference between the square roots of the hydrogen partial pressure in the retentate and

permeate sides at 400 C. ............................................................................................. 54


of the difference between the H2 partial pressure in the retentate and permeate sides at 400 C. .... 54


of the difference between the square roots of the hydrogen partial pressure in the retentate and

permeate sides at 300 C. ............................................................................................. 55


of the difference between the H2 partial pressure in the retentate and permeate sides at 400 C. .... 55

Figure A.7 - H2 flux as a function of the CO concentration in the feed for different H2 trans-menbrane

partial pressure differences at 400 C for the 3.5-4.0 m thick Pd membrane. ............................. 57

Figure A.8 - H2 flux as a function of the CO2 concentration in the feed for different H2 trans-menbrane

partial pressure differences at 400 C for the 3.5-4.0 m thick Pd membrane. ............................. 57

Figure B.1 - Calibration curve for N2. ............................................................................... 61

Figure B.2 - Calibration curve for CO................................................................................ 61

Figure B.3 - Calibration curve for CO2. .............................................................................. 62

Figure C.1 - Linear regression of the hydrogen flux through the Pd-Ag membrane as a function of the

difference between the square roots of the hydrogen partial pressure in the retentate and permeate

sides at 300 C. .......................................................................................................... 63


difference between the hydrogen partial pressure in the retentate and permeate sides at 300 C. .... 64



sides at 400 C. .......................................................................................................... 64




xi



sides at 500 C. .......................................................................................................... 65



Figure C.7 - Comparison of the parity plots for = 0.58 and = 0.50 at 300 C. ........................... 67



Figure D.1 - Effect of temperature on the activity of the 0.5Pt/6CeTiO2 catalyst and on the for

temperatures between 250 and 300 C. ............................................................................. 69

Figure D.2 - Effect of temperature on the activity of the 0.5Pt/6CeTiO2 for temperatures between 250

and 450 C. ............................................................................................................... 70

Figure D.3 - Arrhenius plot for the WGS reaction carried over the 0.5Pt/6CeTiO2 catalyst for the

temperature range between 250 and 300 C and the following inlet gas volume composition: 5% CO,

7.5% CO2, 40% H2O and 35% H2 balanced with N2. .................................................................. 70

Figure D.4 Determination of the apparent WGS reaction order for CO for the 0.5Pt/6CeTiO2 catalyst

at 275 C and 1 bar total pressure. ................................................................................... 71

Figure D.5 - Determination of the apparent WGS reaction order for H2O for the 0.5Pt/6CeTiO2 catalyst


Figure D.6 - Determination of the apparent WGS reaction order for H2 for the 0.5Pt/6CeTiO2 catalyst at

275 C and 1 bar total pressure. ...................................................................................... 72

Figure D.7 - Determination of the apparent WGS reaction order for CO2 for the 0.5Pt/6CeTiO2 catalyst


Figure E.1 - Conversion of CO obtained for the base case at the beginning of each day of the

experimental campaign and for all the other measurements. .................................................. 76

Figure F.1 - Relation between the H2 permeance and the partial pressure of CO. Comparison between

the experimental data and the Sieverts-Langmuir model prediction. ......................................... 78

Figure G.1 - Relative effect of the concentration polarization and poisoning of the Pd-Ag membrane on

the permeating flux of H2 for a feed volumetric composition of 5% CO and 95% H2 and different H2

partial pressure differences in the membrane. .................................................................... 80





xii





xiii

Tables Syllabus

Table 1 Typical WGS inlet stream compositions (vol.%) reported in the literature for WGS tests. ......9

Table 2 Apparent activation energies and reaction orders for the forward WGS reaction. .............. 13

Table 3 Data for different Pd-based membranes from the literature. ...................................... 20

Table 4 - Overview of the operating conditions investigated. .................................................. 25

Table 5 Overview of the operating conditions investigated regarding CO poisoning. ..................... 25

Table 6 Results of error minimization. ............................................................................ 29

Table 7 Results of error minimization in the neighbourhood of . .................................... 29

Table 8 Apparent activation energy and pre-exponential factor for hydrogen permeation through the

dense Pd-Ag membrane used in this work and taken from the literature. .................................... 30

Table 9 - Overview of the operating conditions investigated. .................................................. 33

Table 10 Values of the parameters of the Sierverts-Langmuirs model equation. ........................ 42

Table A.1 Overview of the conditions investigated before the detection of N2 leaks. ................... 52

Table A.2 Results of error minimisation for the 3.5-4.0 m thick Pd membrane. ......................... 56

Table A.3 Results of error minimisation for the 4.0-5.0 m thick Pd membrane. ......................... 56

Table B.1 - Calibrations done for N2, CO and CO2. ................................................................ 59

Table B.2 - Binary mixtures used for the calibration of N2. ..................................................... 59

Table B.3 - Binary mixtures used for the calibration of CO. .................................................... 60

Table B.4 - Binary mixtures used for the calibration of CO2. ................................................... 60

Table D.1 - Apparent activation energy and pre-exponential factor ( ) obtained for the WGS reaction

over the 0.5Pt/6CeTiO2 catalyst. ..................................................................................... 71

Table D.2 - Apparent partial reaction orders obtained for the WGS reaction over the 0.5Pt/6CeTiO2

catalyst at 275 and for a base volumetric composition of 5% CO, 7.5% CO2, 40% H2O and 35% H2

balanced with N2. ........................................................................................................ 73


Introduction 1

1 Introduction

1.1 Framework and Project Presentation

For thousands of years humans have felt the need of using extracorporeal sources of

energy simply to heat themselves, to pump water, to move a vehicle or to keep a television

on in a rainy Saturday night. At first the human race resorted to the burning of wood and

straw, later to the use of the energy of the wind and water, the use of engines based on the

ability to harness the power of steam and for many years fossil fuels have been the source of

energy on which the worldwide society has been relying the most. [1]

With all the environmental problems identified as a consequence of the use of fossil

fuels and also with the still increasing consumption of fossil fuels and the therefore escalating

prices of those, renewable energy sources like hydro energy, solar energy and wind energy

are becoming more important. However, these renewable sources of energy are not enough

to completely take over from the fossil fuels. Hydrogen as an energy carrier has been

commonly thought to play an important role in the future in fuel cells as a substitute for

conventional internal combustion engines and gas turbines because of, for example, its

higher power density and cleaner exhausts. [2]

Nowadays hydrogen is mostly used in petroleum refining processes such as hydrotreating

and hydrocracking and in the petrochemical industry for the production of methanol,

ammonia and hydrocarbon synthesis via the Fischer Tropsch process. There are many routes

for hydrogen production being that the production processes can be categorized into five

types: reforming, electrolysis, nuclear based, photo-catalytic and non-catalytic processes.

Considering the feedstocks used, these hydrogen production processes can be divided into

fossil based processes that use natural gas, coal, methanol and naphtha, and non-fossil based

processes that use water and biomass. [2]

The fossil based routes are the most used for the industrial production of hydrogen,

being that the steam reforming of methane (SRM) is responsible for almost 48% of the

worldwide hydrogen production. The reforming of naphtha/ oil contributes with 30% and the

coal gasification with 18%. [3] A traditional reforming process flow scheme for such a

hydrogen plant is presented in Figure 1.

The traditional reforming process involves firstly a feed treatment in order to remove

the impurities that are poisonous to the reforming and shift catalysts. The most important

process step is the reforming section which can be divided in two sub-steps: the processing of


Introduction 2

the feedstock by reforming or gasification and the water-gas shift (WGS) reaction that

upgrades the carbon monoxide to hydrogen.

Figure 1 - Process flow in a typical fuel processor operating in the steam reforming mode.

Taken from [4]

The basic reactions for the production of hydrogen from natural gas (primarily methane) are

as follows:

Endothermic SRM

CH4 + H2O CO + 3H2 ( = 206 kJmol-1) (1)

Exothermic WGS reaction

CO + H2O CO2+ H2 ( = -41.1 kJmol-1) (2)

The amount of carbon monoxide can still be decreased through catalytic methanation in the

CO elimination step of Figure 1. After methanation other methods such as pressure swing

adsorption (PSA), cryogenic distillation or membrane technology can be used to purify even

more the hydrogen stream. [5]

Although the SRM is the best current option for hydrogen production, in the future the

production of hydrogen may be done through the steam reforming of liquid fuels (e.g. ethanol

and methanol obtained from biomass). [4]

As previously mentioned hydrogen fuel cells may take over the conventional internal

combustion engines in the future because of being more environment friendly and more

efficient. Polymer electrolyte membrane fuel cells (PEMFCs) can generate and deliver electric

power in a wide range that can go from micro to mega-watt and because of that they are

suitable for many different applications at many different scales (from mobile phones to

stationary power stations). Besides that PEMFCs are compact, modular, operate at relative

low temperatures (80-110 C), have high power density, present fast start-up and response

time and have no shielding requirements for personal safety. However, PEMFCs applied to

road vehicles still present some technological limitations associated to water management


Introduction 3

and CO sensitivity of the anode catalyst. One of the main goals at the moment is to reduce

the CO concentration in the H2 stream fed to the fuel cells to a value lower than 0.2 ppm

(ISO, 2008). By doing this it will be possible to avoid CO poisoning of the anode, to reduce the

size and increase the efficiency of the PEMFCs. At the moment, the WGS reaction technology

may be the most promising process to reduce the CO content in the H2 streams and thats why

it is such a hot topic. [4]

1.2 SMR Presentation

The research group Multiphase Reactors (SMR) is a part of the Faculty of Chemical

Engineering at the Technical University of Eindhoven in the Netherlands. The research group

SMR focusses on the fundamentals of the discipline of chemical reaction engineering. The

main area of interest of SMR is the quantitative description of transport phenomena (including

fluid flow) and the interplay with chemical transformations in multiphase chemical reactors.

One of the main goals of SMR is the generation of new knowledge and the development of

new reactor models with improved predictive capability for this industrially important class of

chemical reactors. Through the intended co-operation with other (application oriented)

research groups, both fundamental aspects and those closely related to applications are

studied through concerted action. [6]

1.3 Work Contributes

In this project a highly permeable Pd-based membrane for hydrogen permeation and a

highly active Pt-based catalyst are integrated into a packed bed membrane reactor (PBMR) in

order to analyse the performance of the reactor at different operating conditions:

temperature, gas hourly spacial velocity (GHSV) and carbon/steam ratio. The experimental

data obtained is used to validate both existing 1D and 2D phenomenological models. For the

2D model there is a special attention on the possible verification of the existence of the

concentration polarization effect, which is considered by the model.

1.4 Thesis Organization

This Masters dissertation is divided in 4 chapters, being that the first chapter

encompasses an introduction to the subject of this project as well as a short presentation of

the research group SMR in which all the work was developed and the main goals of the

project. The second chapter is called Context and State of the Art and consists on a

literature review about the current state of all the important subjects addressed.

Chapter 3 is divided in 4 sub-chapters. The sub-chapter 3.1 is called Membrane

characterization and includes a description of the set-up used for the permeation tests as


Introduction 4

well as a list with all the permeation tests performed and the respective results. A discussion

of these results is also included. Sub-chapter 3.2 encompasses a description of the PBMR set-

up and the list of WGS experiments performed on it as well as the respective results and

discussion. Sub-chapter 3.3 includes a short description of both 1D and 2D phenomenological

models developed by another Masters student in parallel with the experimental work reported

in this thesis. A comparison between the experimental results and the phenomenological

models predictions is as well included in order to verify if both models describe adequately

the PBMR system, with special focus on the 2D model in order to verify if indeed there is

concentration polarisation inside the PBMR. A discussion of the results is included.

In the fourth chapter there are included all the conclusions of the work and in the

fifth chapter there are presented all the achieved goals, the limitations of the work,

possibilities for future work and a final appreciation of this dissertation project.


Context and State of the Art 5

2 Context and State of the Art

2.1 History of Hydrogen and the WGS Reaction

Hydrogen has been target of interest by a huge scientific community in the last decades.

However it was way before the 20th or 21st century that hydrogen was subject of research.

Sometimes the first discovery of hydrogen gas is attributed to the Swiss alchemist Philippus

Aureolus Paracelsus in 1520. Paracelsus firstly described a gaseous substance arising as iron

that was dissolved in sulphuric acid. He described this substance as an air which bursts forth

like the wind. [7] In 1671 an English chemist and physicist called Robert Boyle published a

paper called New experiments touching the relation between flame and air in which he

described the reaction between iron filings and diluted acids which results in the formation of

hydrogen. [8] However hydrogen was only identified as a distinct element by British scientist

Henry Cavendish in 1766 after he has separated hydrogen gas by making metallic zinc react

with hydrochloric acid. Cavendish demonstrated to the Royal Society of London that by

applying a spark to hydrogen gas (in the presence of air) it is possible to produce water. This

led him to discover that water is made of hydrogen and oxygen.

In 1783 Jacques Alexander Cesar Charles launched the first hydrogen balloon flight and

the name hydrogen was given to the gas by Antoine Lavoisier in 1788. In 1800 William

Nicholson and Sir Anthony Carlisle discovered the water electrolysis process. Later in 1839 a

Swiss chemist called Christian Friedrich Schoenbein discovered the fuel cell effect which

consists on combining hydrogen and oxygen to produce water and an electric current. This

discovery was later demonstrated by Sir William Grove on a practical scale through the

creation of a gas battery. For this achievement he gained the title of Father of the Fuel

Cell. In the 1920s Rudolf Erren converted the internal combustion engines of trucks, buses

and submarines to use hydrogen or hydrogen mixtures and J.B.S. Haldane introduced the

concept of renewable hydrogen. In 1958 the United States formed the National Aeronautics

and Space Administration (NASA) and currently NASAs space program use the most liquid

hydrogen worldwide, primarily for rocket propulsion and as a fuel for fuel cells. One year

later Francis T. Bacon built the first practical hydrogen-air fuel cell with a power of 5kW.

Later that year Harry Karl Ihrig demonstrated the first fuel cell vehicle: a 20-horsepower

tractor. Hydrogen fuel cells based on Bacons design have been used to generate on-board

electricity, heat and water for astronauts aboard all the space shuttle missions after Apollo

spacecraft.

In 1974 it was formed the International Association for Hydrogen Energy (IAHE) and in

1977 the International Energy Agency (IEA) was established in response the global oil market



disruptions. The activities performed in IEA included the research and development of

hydrogen energy technologies. During the year of 1990 the worlds first solar powered

hydrogen production plant became operational. In 1991 Georgetown University in Washington,

D.C began developing three 3foot Fuel Cell Test Bed Buses as part of their Generation I Bus

Program. Ten years later they finished their Generation II Bus which uses hydrogen from

methanol to power a 100kW fuel cell engine. During the year of 1998 Iceland unveiled a

plan to create the first hydrogen economy by 2030. In 1999 the first European hydrogen

fuelling stations were opened in Hamburg and Munich and two years later Ballard Power

Systems lunched the worlds first volume-produced proton exchange membrane fuel cell

system designed with the aim of being integrated into a wide variety of industrial and

consumer end-product applications. In 2003 U.S.A. announced an investment of $1.2 billion in

a hydrogen fuel initiative to develop the technology for commercially viable hydrogen-

powered fuel cells and in the following year a $350 million investment on hydrogen research

and vehicle demonstration projects was also done. [9]

However, despite all the applications hydrogen has been used for, a simple question

has to be answered: how has hydrogen been produced? As mentioned before, there are many

ways to produce hydrogen being that the most used process is the reforming of methane. This

process can be divided in two steps: the SRM and the WGS. On this thesis the WGS reaction

will be the target of research.

The WGS reaction has been researched for many decades and because of that a vast

amount of knowledge about it has been gathered. Ever since its first industrial application

considerable research regarding reaction catalyst, process configuration, reactor design,

reaction mechanisms and kinetics has been done.

The WGS reaction was observed for the first time in 1780 by Felice Fontana. At the time,

Fontana observed that a combustible gas is produced when steam is passed through a bed of

incandescent coke. On the following century Ludwig Mond developed the process to produce

the so called Mond gas (the product of the reaction of air and steam passed though

coal/coke CO2, CO, H2, N2, etc.), which turned to be the basis for future coal gasification

processes. Mond and his assistant Carl Langer were the first to use the term fuel cells while

performing experiments with the first ever working fuel cell using coal-derived Mond gas.

Their biggest difficulty was to feed pure hydrogen to the Mond battery because of the large

quantities of carbon monoxide present in the Mond gas, which poisoned the Pt electrode. In

order to solve this problem Mond passed Mond gas and steam over finely divided nickel at

400 C, reacting carbon monoxide and steam to produce carbon dioxide and more hydrogen.

After removing CO2 with an alkaline wash, the hydrogen stream could be successfully fed to

the hydrogen cell without poising its electrode. [4] Mond and Langer discovered and reported



for the first time in the literature the WGS reaction in 1888. The first industrial application of

this reaction happened in 1913 for the production of synthesis gas, as a part of the Haber-

Bosch process of ammonia manufacture. This reaction started being considered as a very

important step when it was found that the Fe-based catalyst used in the ammonia synthesis

process was polluted and therefore deactivated by carbon monoxide. This meant that the

carbon monoxide had to be upgraded to hydrogen and carbon dioxide via the WGS reaction.

The WGS reaction was firstly integrated on an industrial scale with the aim to convert the CO

in the syngas produced and at that time the reaction was done in a single stage reactor that

could reduce the CO level to around 10000 ppm (1%). Since the value was still high, a two-

stage system combined with a better catalyst was adopted instead and it resulted in a CO

level lower than 0,5%. The WGS reaction followed the increase of hydrogen demand for the

production of mainly methanol and ammonia. Also the increasing interest in the production of

hydrogen for fuel cells applications required continuous research on the WGS reaction

because of the high purity hydrogen needed for these cells, which are highly sensitive to CO

poisoning. [4,5,10]

2.2 Thermodynamics of the WGS Reaction

The WGS reaction is an equilibrium-limited reaction and since it is exothermic (equation

(2)) the CO conversion and therefore hydrogen production are favoured at lower

temperatures as can be seen in equation (3):

(

) (3)

in which is the equilibrium constant and T the absolute temperature. As the temperature

increases, decreases and consequently the CO conversion at the equilibrium also

decreases, as can be seen in Figure 2. Also lower temperatures are favourable from steam

economys point of view. However, the WGS reaction is kinetics controlled at these

conditions, consequently requiring highly active and stable WGS catalysts. Typically the WGS

reaction is conducted in a two or three-stage converter rather than only one. This

embodiment allows a smaller adiabatic temperature rise and a better steam management

making the process more economical. The first stage is a high temperature converter that

allows a fast CO consumption and the minimization of the catalyst bed volume. The next

stages operate at lower temperatures in order to achieve higher conversions, which are

limited by the reaction equilibrium. [5]



Figure 2 CO equilibrium conversions of a typical reformate stream from a SRM process for

different steam to dry gas (S/G) ratios. Taken from [5]

However, temperature is not the only parameter that affects the CO equilibrium

conversion. As can be seen in Figure 2 the amount of steam that is fed to the WGS reactor

also influences the CO conversion obtained at

he equilibrium, especially for temperatures higher than 150 C. Of course the amount of

steam added to the WGS reactor inlet stream must be decided taking into consideration the

operating conditions, the catalyst capacity for H2O activation, the CO composition desired at

the end of the process and the steam available. The composition of the syngas fed to the WGS

reactor also affects the CO conversion at the equilibrium ( ), as expressed in equation

(4):

( )( )

[ ( )]( ) (4)

being that is the molar fraction of species i at the reactor inlet.

Some of the typical WGS inlet stream compositions used by different authors for WGS tests

are presented in Table 1.



Table 1 Typical WGS inlet stream compositions (vol.%) reported in the literature for WGS

tests.

Component Kalamaras et al. [11]a Roh et al. [12]b Gonzlez et al. [13]b

CO 0.05-9.00 6.5 4.4

CO2 0.00-18.00 7.1 8.7

H2O 3.00-20.00 28.7 29.6

H2 0.00-50.00 42.4 28.0

CH4 - 0.7 0.1

a - Balanced with He. b - Balanced with N2.

2.3 Catalysts for the WGS Reaction

As mentioned before the catalyst plays a very important role in the WGS reaction. Fe-

based catalysts, whether or not combined with Cr, were some of the earliest heterogeneous

catalyst to be used in the industry for the WGS reaction. Maroo et al. [14] reported the

catalytic activity of a Fe-Cr catalyst for the WGS reaction that allowed the attainment of a

CO conversion of 93% at around 380 C. For lower temperatures the CO conversion decreased

very fast meaning that this catalyst is not completely suitable for low temperature WGS. In

another work Maroo et al. [15] reported the performance of another Fe-Cr based WGS

catalysts prepared by co-precipitation and oxi-precipitation. Both catalysts showed good WGS

catalytic activity however, they require high temperatures. These catalysts were also

compared with a commercial WGS catalyst which consisted on a mixture of iron, chromium

and copper oxides. It was noticed that the commercial catalyst was able to catalyse the WGS

reaction at lower temperatures probably because of the presence of copper in its

composition. Mahadevaiah et al. [16] reported the catalytic activity of two new WGS

catalysts: Ce0.67Fe0.33O2- and Ce0.65Fe0.33Pt0.02O2- . It was observed that the platinum catalyst

presents much higher catalytic activity for the WGS reaction than the other one, even at

medium temperatures (300 C), due to the synergistic interaction of the Pt ion with Ce and

Fe ions. The Fe-based catalysts are known to be cheap and stable but, as it has been

concluded, only at high temperatures. The introduction of Cu-based catalysts in the WGS

reaction was a very important step forward because of the higher CO conversion and yield in

the production of H2 that they allowed to obtain at lower temperatures, as seen in one of

Maroos work. [5] However, Cu-based catalysts are sensitive to sulphur and chlorine and so



the syngas needs to be properly cleaned before entering in the WGS reaction stage. Also,

these catalysts normally only operate within a limited temperature range because of

problems with Cu sintering. [5] Zhang et al. [17] showed the sensitivity that Cu-based

catalysts show in the presence of H2S. In their work Zhangs group analysed a series of Fe-

based catalysts for the WGS reaction and characterized them after their exposure to H2S. The

Cu-containing catalysts showed a higher sensitivity to H2S and faster deactivation kinetics

than the Cu-free catalysts. This deactivation is probably due to the fact that the catalyst

surface oxygen is partially replaced by sulphur, which results in pronounced changes in Fe and

Cu coordination environment. [17]

The search for better catalysts that are not only active and stable but also resistant to

impurities, like sulphur, continued and Co-based catalysts were found to fulfil all these

requirements. However, these catalysts only work well at high temperatures and increase the

production of by-products. [5] Au-based catalysts were thought to be promising because of

their high activity at low temperatures. Gamboa-Rosales et al. [18] showed that the Au-

Co3O4/CeO2 bimetalic catalysts allow higher CO conversion and H2 yield than the Co3O4/CeO2

catalysts because of the higher dispersion of gold and reducibility. The gold catalysts also

showed higher activity for lower temperatures. Sakurai et al. [19] verified that some gold

catalysts, like Au (4.0 wt.%)/CeO2, show high WGS activity at temperatures below 250 C.

However, gold catalysts in general were found to deactivate in a reasonable short period of

time under WGS conditions. [5] There have been many explanations to describe de reasons for

gold catalysts deactivation like sintering of the metal particles, irreversible over-reduction

of the ceria support, loss of ceria surface area and blocking of the ceria surface by formation

of surface carbonates and/or formats. However, there isnt still a totally coherent

explanation for this activity reduction because of the many catalyst morphologies and surface

compositions. [5] Pt-based catalysts are very similar to Au-based catalysts, which makes them

very strong candidates as well. [5] Roh et al. [12] reported highly active nanosized (1 wt.%

Pt/CeO2) catalysts for the WGS reaction. During their synthesis a white precursor named

cerium (III) carbonate precipitated and that precipitate was digested for 0-8 h. When there

was no digestion or it was only 2 hours long the Pt catalysts showed lower CO conversion

under medium temperatures than for the cases in which the digestion time was 4 and 8 hours.

For a 4 hours digestion time the Pt catalyst presented the highest CO conversion (80%) and

CO2 selectivity (100%). [12] Jeong et al. [20] tested Pt catalysts over CeO2, ZrO2 and Ce(1-

x)Zr(x)O2 for a single stage WGS reaction. The Pt/CeO2 catalyst presented a CO conversion of

approximately 85% while the Pt/Ce0.8Zr0.2O2 catalyst converted 80% of the CO approximately,

both at the same temperature between 300 and 350 C. The Pt/Ce0.6Zr0.4O2 catalyst

presented a CO conversion slightly lower than 80% for the same temperature while the

Pt/Ce0.4Zr0.6O2 catalyst only originated this conversion at a temperature slightly above 350 C.



The other Pt catalysts with lower cerium content support (20% and no cerium) showed low CO

conversions. Also the CeO2 supported catalyst besides presenting a good stability was the one

that presented the lowest activation energy. This makes Pt/CeO2 a promising catalyst for

single stage WGS reaction. [20] Also, Gonzlez et al. [13] compared the performance of three

different Pt catalysts supported on CeO2, TiO2 and Ce-TiO2. The Pt catalyst supported on Ce-

modified TiO2 support was the one that presented the best activity and better stability at

temperatures higher than 300 C than that of the TiO2 supported one. Gonzlez et al. believe

that the fact that the contact between Pt and Ce in the Pt/Ce-TiO2 catalyst eases the

reducibility of the ceria component in the support at lower temperatures is the cause for the

better activity and stability of this catalyst for the WGS reaction. [13]

In order to decide what catalyst is going to be used some more aspects need to be

considered. Has it is known, the presence of WGS products (H2 and CO2) is disadvantageous

since they inhibit the WGS catalysts thus lowering the reaction rate. This inhibition effect

depends not only on the nature of the catalyst thats used but also on the temperature range

at which the reaction occurs. Consequently the reaction effectiveness can be highly improved

when membrane reactors (MRs) are used instead of traditional reactors (TRs). For the case of

using MRs for the WGS reaction two situations are possible depending on the nature of the

membrane used. If a hydrogen perm-selective membrane is used, the concentration of CO2 in

the reaction medium will be higher, which affects the reaction rate. If a CO2 perm-selective

membrane is used, then the hydrogen concentration in the reactor medium will be high thus

affecting also the reaction rate. For some Fe-based catalysts the presence of hydrogen at high

concentrations is adverse since it over-reduces the magnetite active phase.

As can be seen in Table 2 Cu-based catalysts performance is slightly more affected in

the presence of WGS products than Au or Pt catalysts, for similar temperatures, pressures and

feed composition. By making a comparison between Pt and Au catalysts it may be fair to say

that Pt-based WGS catalysts are one step ahead since they present lower inhibition by WGS

products than Au-based catalysts (Table 2) and also, because the science of Au catalysis is

relatively new and so theres a bigger know-how about Pt catalysts. Therefore a Pt-based

catalyst is used in this project for the WGS reaction carried out inside a MR.

2.4 Mechanisms and Kinetics

In this section a small review on the WGS reaction kinetics and mechanisms is done.

The reaction rate for the WGS reaction is normally written as follows:

(5)



(

) (6)

(7)

where is the experimental reaction rate, is the forward reaction rate, is the pre-

exponential factor, is the approach to equilibrium, is the partial pressure of component

, is the apparent activation energy and is the gas constant.

In Table 2 there are presented some important literature values of kinetic parameters

obtained for some of the most relevant WGS catalysts. In terms of apparent activation

energy, a comparison between both Fe3O4/Cr2O3 catalyst and 1% Pt/Al2O3 catalyst (higher

temperatures) can be done being thus possible to notice that even without WGS products in

the feed stream, the Fe-based catalyst presented a higher apparent activation energy than

the Pt-based catalyst. This suggests that the reaction mechanism or the rate-determining step

for the Fe catalyst may be different from that of the other catalysts. By making a comparison

between the Al2O3 supported catalysts and the CeO2 supported catalysts in terms of apparent

activation energy, it is possible to conclude that overall the CeO2 supported catalysts present

lower apparent activation energy. In particular, the 8% CuO/15% CeO2/Al2O3 catalyst is the

one that presents the lowest apparent activation energy. This might be due to the increase of

the reducibility of the surface oxygen in the ceria support, which probably results from the

addition of Cu. [5]

Normally the reaction rate data are fitted to a power-law with the following form:

(8)

where are the forward reaction orders and is the forward reaction rate constant. By

analysing Table 2 once again it can be observed that for the Pt-based catalysts all the

apparent reaction orders except the H2O reaction order are very similar. Also, the apparent

activation energies are quite close. In terms H2O apparent reaction orders for Pt catalysts,

the alumina-supported catalysts present values close to 1 while the ceria-supported ones

present values near to 0.5. This difference suggests that different reaction mechanisms

and/or different sites for H2O activation may exist for these materials. Also by analysing the

H2O apparent reaction orders for both Au-based catalysts in Table 2 it can be concluded that

the 4.5 wt% Au/CeO2 catalyst, for which the H2O concentration at the inlet stream was higher,

presents a H2O apparent reaction order higher than the 2.6 wt% Au/CeO2 catalyst. This can be

explained by the occurrence of H2O dissociation on the catalyst surface, where OH groups

react with hydrogen to produce water. This means that the WGS reaction is sensitive towards

the partial pressure of water in the feed stream. Also, competitive adsorption between H2O

and H2 may happen. [5]



Table 2 Apparent activation energies and reaction orders for the forward WGS reaction.

Catalyst Operating

conditionsa

Ea

(kJmol-1)

Reaction orderb Reference

H2O CO H2 CO2

1% Pt/Al2O3 1 atm, 285 C 68 1.0

(10-46%)

0.06

(5-25%)

-0.44

(25-60%)

-0.09

(5-30%)

[21]

1% Pt/Al2O3 1 atm, 315 C 84 1.1

(10-46%)

0.1

(5-25%)

-0.44

(25-60%)

-0.07

(5-30%)

[21]

1% Pt/CeO2 1 atm, 200 C 75 0.44

(10-46%)

-0.03

(5-25%)

-0.38

(25-60%)

-0.09

(5-30%)

[21]

2% Pt/CeO2-

ZrO2

1.3 bar, 210-

240 C

71 0.67 0.07 -0.57 -0.16 [22]

2% Pt-

1%Re/CeO2-

ZrO2

1.3 bar, 210-

240 C

71 0.85 -0.05 -0.32 -0.05 [22]

4.5 wt%

Au/CeO2

1bar, 180 C - 1.0

(5-20 kPa)

1.0

(2-5 kPa)

-0.7

(50-78 kPa)

-0.5

(5-20 kPa)

[23]

2.6 wt%

Au/CeO2c

1bar, 180 C 40 0.5

(0.7-10 kPa)

0.5

(0.2-2 kPa)

-0.5

(3.2-75 kPa)

-0.5

(1.2-3.4 kPa)

[24]

40%

CuO/ZnO/Al2O

3

1 bar, 190 C 79 0.8

(10-46%)

0.8

(5-25%)

-0.9

(25-60%)

-0.9

(5-30%)

[25]

8% CuO/CeO2 1 bar, 240 C 56 0.4

(10-46%)

0.9

(5-25%)

-0.6

(25-60%)

-0.6

(5-30%)

[25]

8% CuO/Al2O3 1 bar, 200 C 62 0.8

(10-46%)

0.9

(5-25%)

-0.8

(25-60%)

-0.7

(5-30%)

[25]

8% CuO/15%

CeO2/Al2O3

1 bar, 200 C 32 0.6

(10-46%)

0.7

(5-25%)

-0.6

(25-60%)

-0.6

(5-30%)

[25]

Fe3O4/Cr2O3 1 bar, 450 C 118 0.0

(20-75%)

1.0

(10-40%)

- - [26]

a - Temperature and total pressure at which the reaction order experiments were carried out. b - The values between brackets are the ranges of concentrations for each species in the feed,

or their partial pressures. c - The reaction orders for H2O and CO were obtained in the absence of H2 and CO2 in the feed

stream.



In terms of CO apparent reaction order, the values for Pt catalysts are close to zero

(sometimes positive, sometimes negative). Since CO highly adsorbs on the Pt surface its

coverage is close to saturation and thus, an increase in the CO partial pressure normally has

no effect on the reaction rate or can even reduce it due to surface blocking, even at high

pressures and low temperatures. For Au-based catalysts it can be seen that an increase in the

CO partial pressure increases its apparent reaction order. This happens because CO

adsorption on Au is very weak and so only a low coverage of the surface is obtained for lower

pressures. [5]

Regarding the H2 apparent reaction order for Pt catalysts it can be observed that H2

highly adsorbs on Pt thus inhibiting the WGS reaction. This may be explained by the possibility

that after CO has achieved the coverage saturation, the free Pt sites available for H2O

activation may be occupied by atomic hydrogen thus inhibiting the forward WGS reaction. The

same situation can be observed for the case of Au catalysts. Finally, for the case of CO2

partial reaction order the low values in Table 2 for Pt catalysts may be due to the weak

interaction between CO2 and Pt. On the other hand, for Au catalysts the inhibition effect is

much more significant. This may be explained by the blocking of ceria surface sites and/or

increase of the reverse reaction due to a higher amount of carbonate species adsorbed at the

active sites. [5]

As mentioned before, Pt catalysts present different apparent reaction orders

depending on, for example, the material used in the support. This suggests that there are

different reaction mechanisms or rate-limiting steps for the WGS reaction over this kind of

materials. Until now two main reaction mechanisms have been proposed: the regenerative

mechanism and the Langmuir-Hinshelwood mechanism.

The regenerative mechanism, sometimes called oxidation-reduction cycle,

encompasses first the dissociation of water on the catalyst surface producing H2 and oxidizing

the respective active site (*). In order to complete the cycle the reduction of the oxidized

site is promoted by a CO molecule yielding CO2. This two steps mechanism can be presented

as follows [5]:

H2O + (*) (9)

CO + (O) CO2 + (*) (10)

Considering that equation (10) is the rate-limiting step, one of the earliest WGS reaction rate

expression was proposed as follows:

(

)

(

)

(11)



where is the forward reaction rate constant and is the rate constant of the reverse

reaction.

The Langmuir-Hinshelwood mechanism involves 6 steps:

Dissociative adsorption of water to form reactive hydroxyl groups;

Adsorption of CO;

Combination of the adsorbed CO with the reactive hydroxyl groups to form an

intermediate structure that normally is a formate and/or carbonate;

Decomposition of the intermediate structure into CO2 and H2.

H2O + 2(*) H* + OH*(12)

CO + (*) CO* (13)

OH* + CO* HCOO* + (*) (14)

HCOO* + (*) CO2(*) + H(*) (15)

CO2* CO2 + (*) (16)

2H* H2 + (*) (17)

A kinetic expression (equation (18)) for the WGS reaction rate over a Cu-based catalyst at low

temperatures was proposed considering that the surface reaction between molecularly

adsorbed reactants to form a formate intermediate and atomically adsorbed hydrogen is the

rate-limiting step.

(

)

(

)

(18)

Where is the forward reaction rate constant and is the equilibrium adsorption constant of

species . [5]

2.5 Hydrogen Purification

The purification of the hydrogen stream that comes either directly from the reactor or

from the catalytic methanation process is a very important step towards the achievement of a

final high pure hydrogen stream. There are different methods to purify hydrogen like PSA,

cryogenic distillation, or membrane separation. Unlike traditional processes, which produce a

stream with medium purity (94-97%) of hydrogen, the PSA process allows the attainment of

99,9% purity hydrogen. The PSA process has been used since the 1980s in almost all hydrogen

plants not only because of the higher purification of hydrogen but also because it requires



less unit operations and consequently is less complex from the operational point of view. The

PSA-based process needs only a high temperature WGS stage while the traditional process

requires not only the high temperature stage but also the low temperature one. The PSA-

based process is also advantageous because of the lower steam to carbon ratios that it

requires and because it produces a hydrogen stream completely free of methane. [5] In Figure

3 there are presented the process schemes for both traditional and PSA-based hydrogen

production and purification.

Besides the PSA process there is an even more promising purification process:

membrane separation. The membrane separation process is based on the selective

permeation of hydrogen through the membrane. Separation membranes have potential to be

long-lasting and cheap what makes them highly attractive. The combination of membrane

separation and the WGS reaction (as can be seen in Figure 3) became a very appealing subject

because for the case of equilibrium-limited reactions, the continuous removal of hydrogen or

carbon dioxide from the reaction medium, depending on the type of membrane used, allows

the shifting of the reaction equilibrium towards the formation of products, in other words

higher conversions. [5, 27]

Figure 3 - Process schemes for the hydrogen production and purification; (a) traditional

process considering an absorption and catalytic approach for hydrogen purification (b) PSA-

based hydrogen purification. Taken from [5]



Figure 4 - Hydrogen production and purification based on the WGS MR unit. Taken from [5]

2.6 Membrane Reactors

A MR is a device for simultaneously carrying out a reaction and a membrane-based

separation in the same physical device. Taken from [5] MRs present consequently many

advantages when compared to TRs, being that the main ones are:

Conversion enhancement of equilibrium-limited reactions;

Enhancement of the hydrogen yield and hydrogen selectivity (in the case of hydrogen

production);

Attainment of the same performance obtained in the TR at milder operating

conditions, meaning that it is possible to reduce material costs and because of the

lower temperatures used new heat integration strategies must be adopted;

Achievement of better performance at the same operating conditions as in the TR;

Reduced capital costs due to the combination of reaction and separation in only one

system. [5]

There are three main types of membranes: organic membranes, inorganic membranes

and organic/inorganic hybrid membranes. Normally inorganic membranes present several

advantages over the organic ones, like superior stability and good chemical and mechanical

resistance at temperatures above 100 C. [5] The inorganic membranes, in particular, can

either be dense or porous, made from metals, carbon, ceramics or glass. Also inorganic

membranes for MRs can be inert or catalytically active. Pd membranes and its alloys with Ni,

Ru or Ag are dense inorganic membranes while alumina membranes, silica membranes, titania

membranes, glass membranes and stainless steel membranes are porous membranes. Usually

dense membranes present higher selectivities for a specific component than porous

membranes. However, permeability is also a very important factor to have in consideration.



As expected dense membranes present lower trans-membrane fluxes than porous

membranes.[2] Considering now the case of the WGS reaction and the attempt to produce

hydrogen enough purified to be used in PEMFCs, it is possible to conclude that it is preferable

to use H2 perm-selective membranes in order to isolate H2 instead of using CO2 perm-selective

membranes and having hydrogen mixed with steam and some unreacted CO. Also, since the

purity of the hydrogen obtained is more important than its quantity, it is preferable to choose

a high hydrogen selectivity membrane (dense inorganic membranes).

2.7 Dense H2 Perm-Selective Membranes for Membrane Reactors

A new challenge that is encountered with MRs is the membrane itself. There are

different types of membranes and the membrane to be used needs to be adjusted so that it

can be used in a MR. Pd-based membranes are the most promising for the WGS MR technology

because of the very high hydrogen selectivity that they present. This means that the

permeate (the stream that goes through the membrane) contains almost pure hydrogen as can

be seen in Figure 5, while the retentate contains all the substances that dont go through the

membrane.

However, Pd membranes are sensitive to poisoning and can suffer some embrittlement

in the presence of H2 at low temperatures. When combined with other metals, like Ag or Cu

Figure 5 Composition of the outlet stream of a MR and a traditional process with a typical

composition of syngas coming out of a reformer and a H2O/CO molar ratio of 1. Taken from

[27]



among others, Pd-based membranes may get less susceptible to these effects. [28] Also it has

been shown by Tosti et al. [29] that self-supported dense and thin wall Pd-Ag 23 wt% tubular

membranes with finger-like configuration have high durability and reliability since complete

hydrogen selectivity and none failure were observed after at least one year of thermal and

hydrogenation cycles. These Pd-Ag 23 wt% membranes were also shown to be highly

permeable to hydrogen. These characteristics all together with the reduced costs make this

technology ready for being used in the production of highly pure hydrogen in energetic and

industrial applications (PEMFCs for example). [29]

Normally it is considered that the permeability of hydrogen through a Pd-based

membrane can be described by a solution-diffusion mechanism and the trans-membrane flux

can be described by the following equation:

(

) (19)

where is the hydrogen flux, is the permeability of the membrane, is the thickness

of the membrane, and are the partial pressures of H2 in the retentate

and in the permeate respectively and is the pressure exponent. The ratio

is normally

termed permeance or pressure normalized flux. The pressure exponent varies between 0.5

and 1. This exponent is 0.5 (Sieverts law) when the diffusion of atomic hydrogen through the

metallic lattice of the membrane is the limiting step. The dependency on temperature of the

permeability of the membrane is described below:

(

) (20)

where is the pre-exponential factor and is the apparent activation energy of the

Pd membrane.

In Table 3 a literature review on different Pd-based membranes that have been

reported over the years is presented. Some parameters such as membrane thickness,

hydrogen flux through the membrane, hydrogen permeability, ideal H2/N2 selectivity and

apparent activation energy are compared.

By analysing Table 3 it can be concluded that thicker Pd-Ag membranes in general

present a better combination of permeability and selectivity of hydrogen, mainly because of

the infinite H2/N2 ideal selectivity, which is very important because of the desired purity for

the hydrogen used in PEMFCs. However they are too thick to be industrially implemented

because of the high costs that its implementation at a high scale would involve. Although

thinner membranes are very promising in terms of permeability and also because of their low

thickness and therefore potential for future industrial implementation, they still are quite

limited in terms of selectivity and durability. In this project thin Pd membranes were initially



used because of the above mentioned potential. However, due to the problems previously

stated a Pd-Ag membrane with a 50 m thickness was finally selected.

Table 3 Data for different Pd-based membranes from the literature.

Membrane (C) (kPa)

(m)

Permeability to H2

(molm-1s-1Pa-0.5)

Ideal

Selectivity

H2/N2

(kJmol-1)

Reference

Pd-Ag 200-300 10-150 50 a 10.72 [4,30]

Pd-Ag 350-400 100-400 61 11,24 [29]

Pd-Ag 352 800 84 a 2.92 [31]

Pd-Cu-Y 400 - 2.0 >10000 - [32]

Pd-Cu-Mo 400 - 2.0 >10000 29.0 [32]

Pd-Au 400 - 2.0 >10000 - [32]

a - Calculated value.

2.8 The WGS Reaction in Packed Bed Membranes Reactors

There are two main types of MRs: PBMR and FBMR). In this thesis the focus is on the

analysis of a PBMR.

Figure 6 Scheme of a PBMR for the WGS reaction. Adapted from [10]

With the membrane integrated in the packed bed reactor only one process unit is

needed for the WGS. Normally Pd-based membranes are used, as mentioned before, so that a

pure hydrogen stream can be produced. After being inserted in the reactor the Pd-based

membrane is filled with a WGS catalyst (if the permeation occurs from the inside to the



outside of the membrane). At increased pressures it is possible to achieve almost complete

conversion of CO. A particularity of PBMR is that there are mass transfer limitations from the

catalyst bed to the membrane surface (concentration polarization) unless a sufficiently small

tube diameter is used. However, decreasing the tube diameter implies using smaller catalyst

particles whose minimum size is restricted by the pressure drop restrictions. [10, 33]

There are 4 important parameters that have always to be considered in a PBMR because

of their influence on the CO conversion and H2 recovery: operating temperature, retentate

pressure, H2O/CO ratio and the gas hourly space velocity.

Mendes et al. [34] showed that by performing the WGS reaction in a PBMR the equilibrium

can be shifted resulting in a higher CO conversion. Also, for increasing temperatures it was

verified that the CO conversion decreased while the H2 recovery kept increasing. This happens

because the permeability of the Pd membrane increases with increasing temperatures and on

the other hand, since the WGS reaction is exothermic higher temperatures dont favour the

CO conversion.

Mendes group also showed that for increasing GHSVs the CO conversion decreases,

especially for lower temperatures. The hydrogen recovery is highly affected as well.

Regarding the pressure influence it was verified that for higher retentate pressures both CO

conversion and H2 recovery increase since the driving force for hydrogen permeation through

the membrane is higher and, consequently the equilibrium is more shifted towards higher

conversions. [34]

In terms of H2O/CO ratio Mendes group analysed the effect of both CO and steam inlet

concentration on the performance of the PBMR. When the CO content is increased while

keeping the steam concentration constant both CO conversion and H2 recovery are negatively

affected. This happens because higher concentrations of CO inhibit more the H2 permeability

of the Pd membrane and in some cases also because the catalyst is negatively affected

because of the negative partial order with respect to CO. On the other hand by increasing the

steam amount while keeping the CO content constant both CO conversion and H2 recovery are

enhanced. This can be explained by the fact that for some catalysts the amount of steam fed

to the reactor is crucial for their performance (the Pt catalysts presented on Table 2 present

this behaviour). Also the Le Chateliers principle was verified for conversions beyond the

thermodynamic equilibrium meaning that higher steam concentrations result in higher H2

production. [34] Also, higher amounts of steam allow avoiding carbon deposition on the WGS

catalyst.



2.9 Modeling Studies on the WGS Reaction Carried in Packed Bed

Membrane Reactors

In the last decades there has been a huge effort on trying to simulate the WGS

reaction in PBMRs as realistic as possible. There have been reported many 1D and 2D models

for PBMRs considering many assumptions.

For 1D models normally the following assumptions are made:

Steady-state and isothermal operation;

Axially dispersed plug-flow pattern in the retentate side with pressure drop described

by Erguns equation;

Radial temperature and concentration gradients are ignored;

Ideal plug flow pattern in the permeate side without any pressure drops;

Ideal gas behavior. [30, 33]

In the case of 2D models generally the following assumptions are comprised:

The mass and energy transport in the gas phase is described as convective flow with

axial and radial dispersion;

The particle size is sufficiently small so that it can be considered that intra-particle

mass and heat transfer limitations as well as external mass and heat transfer

limitations from the gas bulk to the catalyst surface can be neglected;

Homogenous gas phase reactions are ignored due to the relatively low temperatures;

The gas bulk is described as an ideal Newtonian fluid. [2, 33]

Marn et al. [35] concluded that intra-particle mass transfer limitations are not

negligible and so generalized Thiele modulus, apparent kinetic parameters or empirical fitting

of external efficiency were proposed.

In this thesis project the validation of both 1D and 2D existing phenomenological models is

done being that the main focus is on the validation of the radial dispersion, as known as

concentration polarisation effect, considered in the 2D model, since until now it has seldom

been done. Tiemersma et al. [10] reported a 2D model to study the performance of the

autothermal reforming (ATR) of methane in a PBMR. This model included the assumption of

the existence of the concentration polarization effect however, it was not validated with

experimental data for the ATR of methane in a PBMR.


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